Process for the gas-phase (co-)polymerisation of olefins in a fluidised bed reactor

ABSTRACT

The present invention relates to a process for the gas-phase (co-)polymerisation of olefins in a fluidised bed reactor using a Ziegler-Natta type catalyst characterised in that the polymerisation is performed in the presence of a monohalogenated hydrocarbon compound.

The present invention relates to a process for the (co-)polymerisationof olefins using a Ziegler-Natta type catalyst, in particular to aprocess for the gas-phase (co-)polymerisation of olefins in a fluidisedbed reactor using a Ziegler-Natta type catalyst.

The present invention also relates to a process for increasing thepolymerisation activity of a Ziegler-Natta type catalyst during the(co-)polymerisation of olefins using said Ziegler-Natta type catalyst,in particular during the gas-phase (co-)polymerisation of olefins in afluidised bed reactor using said Ziegler-Natta type catalyst. Thepresent invention further relates to a process for improving theprocessability of the polymers obtained during the (co-)polymerisationof olefins using said Ziegler-Natta type catalyst, in particular duringthe gas-phase (co-)polymerisation of olefins in a fluidised bed reactorusing said Ziegler-Natta type catalyst.

Processes for the co-polymerisation of olefins in the gas phase are wellknown in the art. Such processes can be conducted for example byintroducing the gaseous monomer and comonomer into a stirred and/orfluidised bed comprising polyolefin and a catalyst for polymerisation.

In the gas fluidised bed polymerisation of olefins, the polymerisationis conducted in a fluidised bed reactor wherein a bed of polymerparticles is maintained in a fluidised state by means of an ascendinggas stream comprising the gaseous reaction monomer. The start-up of sucha polymerisation generally employs a bed of polymer particles similar tothe polymer, which it is desired to manufacture. During the course ofpolymerisation, fresh polymer is generated by the catalyticpolymerisation of the monomer, and polymer product is withdrawn tomaintain the bed at more or less constant volume. An industriallyfavoured process employs a fluidisation grid to distribute thefluidising gas to the bed, and to act as a support for the bed when thesupply of gas is cut off. The polymer produced is generally withdrawnfrom the reactor via a discharge conduit arranged in the lower portionof the reactor, near the fluidisation grid. The fluidised bed consistsin a bed of growing polymer particles. This bed is maintained in afluidised condition by the continuous upward flow from the base of thereactor of a fluidising gas.

The polymerisation of olefins is an exothermic reaction and it istherefore necessary to provide means to cool the bed to remove the heatof polymerisation. In the absence of such cooling the bed would increasein temperature and, for example, the catalyst would become inactive orthe bed would begin to melt. In the fluidised bed polymerisation ofolefins, the preferred method for removing the heat of polymerisation isby supplying to the polymerisation reactor a gas, the fluidising gas,which is at a temperature lower than the desired polymerisationtemperature, passing the gas through the fluidised bed to conduct awaythe heat of polymerisation, removing the gas from the reactor andcooling it by passage through an external heat exchanger, and recyclingit to the bed. The temperature of the recycle gas can be adjusted in theheat exchanger to maintain the fluidised bed at the desiredpolymerisation temperature. In this method of polymerising alphaolefins, the recycle gas generally comprises the monomer and comonomerolefins, optionally together with, for example, an inert diluent gassuch as nitrogen or a gaseous chain transfer agent such as hydrogen.Thus, the recycle gas is used to supply the monomer to the bed, tofluidise the bed, and to maintain the bed at the desired temperature.Monomers consumed by the polymerisation reaction are normally replacedby adding make up gas or liquid to the polymerisation zone or reactionloop.

It is well known that Ziegler-Natta type catalysts can advantageously beused for the (co-)polymerisation of olefins, particularly in slurryprocesses as well as in gas phase processes.

In the course of their research developments, the applicants have nowfound a new method for increasing up to an unexpected level thepolymerisation activity of a Ziegler-Natta type catalyst during the(co-)polymerisation of olefins using said Ziegler-Natta type catalyst,in particular during the gas-phase (co-)polymerisation of olefins in afluidised bed reactor using said Ziegler-Natta type catalyst, especiallyduring the gas-phase (co-)polymerisation of ethylene in a fluidised bedreactor using said Ziegler-Natta type catalyst. What is also unexpectedfrom these huge activity increases is that the present method isabsolutely not detrimental and rather beneficial to a normal andefficient process behaviour of industrial plants; in this respect, theApplicants have found that their method can be successfully applied forincreasing plant throughput while avoiding the usual fouling problemsthe man skilled in the art would expect to face at these highactivities.

Simultaneously, the applicants have found that this new method allows toimprove the processability of the produced polymers obtained during thecopolymerisation of olefins using a Ziegler-Natta type catalyst, inparticular during the gas-phase copolymerisation of olefins in afluidised bed reactor using a Ziegler-Natta type catalyst, especiallyduring the gas-phase copolymerisation of ethylene with another olefin ina fluidised bed reactor using said Ziegler-Natta type catalyst.

This method is especially valuable for the industrial plants that willbe now able by keeping their actual Ziegler-Natta type catalyst toincrease significantly their polymer production.

In accordance with the present invention, there has now been found aprocess for the gas-phase (co-)polymerisation of olefins in a fluidisedbed reactor using a Ziegler-Natta type catalyst, said process comprisingthe addition into the reactor of an organoaluminium cocatalyst and of amonohalogenated hydrocarbon compound, wherein the molar ratio of themonohalogenated hydrocarbon compound to the cocatalyst is comprisedbetween 0.02 and 0.2, preferably between 0.02 and 0.15.

The organoaluminium cocatalyst and/or the monohalogenated hydrocarboncompound can be added at any location of the fluidised bedpolymerisation system, e.g. in the reactor itself, below thefluidisation grid or above the grid in the fluidised bed, above thefluidised bed, in the powder disengagement zone of the reactor (alsonamed velocity reduction zone), anywhere in the reaction loop or recycleline, in the fines recycle line (when a cyclone is used). According toan embodiment of the present invention, the organoaluminium cocatalystand/or the monohalogenated hydrocarbon compound are added directly intothe fines recycle line (when a cyclone is used), or directly into thepolymerisation zone, more preferably directly into the fluidised bed,ideally in the lower part of the bed (below half bed height). For thepurposes of the present invention and appended claims, thepolymerisation zone means the reaction zone consisting of the fluidisedbed itself, and in the region above the fluidised bed which consists ofthe powder disengagement zone and/or the velocity reduction zone.According to another embodiment of the present invention, theorganoaluminium cocatalyst and/or the monohalogenated hydrocarboncompound are added at at least two different locations of the fluidisedbed polymerisation system. It is preferred according to the presentinvention that the organoaluminium cocatalyst and/or the monohalogenatedhydrocarbon compound are not added in admixture with the catalyst.According to another embodiment, the organoaluminium cocatalyst and/orthe monohalogenated hydrocarbon compound are added into the fluidisedbed polymerisation system through the well-known BP high productivitynozzles, which protrude through the fluidisation grid directly into thefluidised bed (see e.g. WO9428032, the content of which is incorporatedhereby).

The monohalogenated hydrocarbon compound may be a chlorinated orbrominated hydrocarbon. It may be a monohalogenated hydrocarbon ofgeneral formula R—X in which R denotes an alkyl group containing from 1to 10, preferably from 2 to 7 carbon atoms, an aralkyl or aryl groupcontaining from 6 to 14, preferably from 6 to 10 carbon atoms, and Xdenotes a halogen atom such as chlorine or bromine. Preferably, themonohalogenated hydrocarbon compound is chosen amongst methylenechloride, ethyl chloride, propyl chloride, butyl chloride, pentylchloride, hexyl chloride and heptyl chloride. Butyl chlorides are morepreferred, n-butyl chloride being the most preferred monohalogenatedhydrocarbon compound.

According to a preferred embodiment of the present invention, theinvention monohalogenated hydrocarbon compound is diluted in aconventional diluent. Suitable diluents include aromatic, paraffin andcycloparaffin compounds. The diluents are preferably selected from amongbenzene, toluene, xylene, cyclohexane, fuel oil, isobutane, pentane,kerosene and mixtures thereof for instance. When a diluent is used, theinvention monohalogenated hydrocarbon compound is preferably present inan amount comprised between 0.001 and 2 mole/l of diluent, preferablybetween 0.005 and 1 mole/l of diluent. Said diluent is preferablybutane, pentane or hexane. The process according to the presentinvention is particularly suitable for the manufacture of polymers in acontinuous gas fluidised bed process.

In an advantageous embodiment of this invention, the polymer is apolyolefin preferably copolymers of ethylene and/or propylene and/orbutene. Preferred alpha-olefins used in combination with ethylene and/orpropylene and/or butene in the process of the present invention arethose having from 4 to 8 carbon atoms. However, small quantities ofalpha olefins having more than 8 carbon atoms, for example 9 to 40carbon atoms (e.g. a conjugated diene), can be employed if desired. Thusit is possible to produce copolymers of ethylene and/or propylene and/orbutene with one or more C4-C8 alpha-olefins. The preferred alpha-olefinsare but-1-ene, pent-1-ene, hex-1-ene, 4-methylpent-1-ene, oct-1-ene andbutadiene. Examples of higher olefins that can be copolymerised with theprimary ethylene and/or propylene monomer, or as partial replacement forthe C4-C8 monomer are dec-1-ene and ethylidene norbornene. According toa preferred embodiment, the process of the present invention preferablyapplies to the manufacture of polyolefins in the gas phase by thecopolymerisation of ethylene with but-1-ene and/or hex-1-ene and/or4-methylpent-1-ene.

The process according to the present invention may particularly be usedto prepare a wide variety of polymer products for example linear lowdensity polyethylene (LLDPE) based on copolymers of ethylene withbut-1-ene, 4-methylpent-1-ene or hex-1-ene and high density polyethylene(HDPE) which can be for example copolymers of ethylene with a smallportion of higher alpha olefin, for example, but-1-ene, pent-1-ene,hex-1-ene or 4-methylpent-1-ene.

The process is particularly suitable for polymerising olefins at anabsolute pressure of between 0.5 and 6 MPa and at a temperature ofbetween 30° C. and 130° C. For example for LLDPE production thetemperature is suitably in the range 75-105° C. and for HDPE thetemperature is typically 80-120° C. depending on the activity of thecatalyst used and the polymer properties desired.

The polymerisation is preferably carried out continuously in a verticalfluidised bed reactor according to techniques known in themselves and inequipment such as that described in French Patent Application 0004757(filing number), European patent application EP-0 855 411, French PatentNo. 2,207,145 or French Patent No. 2,335,526.

The process of the invention is particularly well suited toindustrial-scale reactors of very large size.

In one embodiment the reactor used in the present invention is capableof producing greater than 300 kg/h to about 80,000 kg/h or higher ofpolymer, preferably greater than 10,000 kg/h.

The polymerisation reaction is carried out in the presence of aZiegler-Natta type catalyst.

Examples of Ziegler-Natta type catalysts according to the presentinvention are typically those consisting of a solid catalyst essentiallycomprising a compound of a transition metal and of a cocatalystcomprising an organic compound of a metal (i.e. an organometalliccompound, for example an alkylaluminium compound). These high-activityZiegler-Natta type catalyst systems have already been known for a numberof years and are capable of producing large quantities of polymer in arelatively short time, and thus make it possible to avoid a step ofremoving catalyst residues from the polymer. These high-activitycatalyst systems generally comprise a solid catalyst consistingessentially of transition metal complexes, magnesium complexes andhalogen containing complexes. Examples thereof can be found, e.g. inU.S. Pat. No. 4,260,709, EP0598094, EP0099774 and EP0175532. The processis also particularly suitable for use with Ziegler catalysts supportedon silica, e.g. in WO9309147, WO9513873, WO9534380, WO9905187 andEP998503. For the purpose of the present description and appendedclaims, Ziegler-Natta type catalysts specifically exclude themetallocene catalysts.

According to a preferred embodiment of the present invention theZiegler-Natta type catalyst consists of a catalyst precursor and of acocatalyst, said catalyst precursor comprising a catalyst carriermaterial, an alkylmagnesium compound, a transition metal compound ofGroups 4 or 5 of the Periodic table of the elements, and an optionalelectron donor.

The catalyst carrier materials that can be used in the present inventionare solid, porous carrier materials such as e.g. silica, alumina andcombinations thereof. They are preferably amorphous in form. Thesecarriers may be in the form of particles having a particle size of fromabout 0.1 micron to about 250 microns, preferably from 10 to about 200microns, and most preferably from about 10 to about 80 microns. Thepreferred carrier is silica, preferably silica in the form of sphericalparticles e.g. spray dried silica.

The internal porosity of these carriers may be larger than 0.2 cm³/g,e.g. larger than about 0.6 cm³/g. The specific surface area of thesecarriers is preferably at least 3 m²/g, preferably at least about 50m²/g, and more preferably from, e.g. about 150 to about 1500 m²/g. It isdesirable to remove physically bound water from the carrier materialprior to contacting this material with water-reactive magnesiumcompounds. This water removal may be accomplished by heating the carriermaterial to a temperature from about 100° C. to an upper limit oftemperature represented by the temperature at which change of state orsintering occurs. A suitable range of temperatures may, thus, be fromabout 100° C. to about 850° C. Preferably, said temperature is comprisedbetween 500° C. and 800° C.

Silanol groups represented by a presence of Si—OH groups in the carrierare present when the carrier is contacted with water-reactive magnesiumcompounds in accordance with the present invention. These Si—OH groupsare usually present at about 0.3 to about 1.2 mmoles of OH groups pergram of carrier, preferably at about 0.3 to about 0.7 mmoles of OHgroups per gram of carrier. Excess OH groups present in the carrier maybe removed by heating the carrier for a sufficient time at a sufficienttemperature to accomplish the desired removal. For example, the silicacarrier, prior to the use thereof in the first catalyst synthesis stephas been dehydrated by fluidising it with nitrogen or air and heating atleast at about 600° C. for at least about 5 hours to achieve a surfacehydroxyl group concentration of less than about 0.7 mmoles per gram(mmoles/g).

The surface hydroxyl concentration (OH) of silica may be determinedaccording to J. B. Peri and A. L. Hensley, Jr., J. Phys. Chem., 72(8),2926 (1968).

The silica of the most preferred embodiment is a material marketed underthe tradename of ES70 by Crosfield and having a surface area of 280 m²/gand a pore volume of 1.6 ml/g. Another preferred silica is a highsurface area, amorphous silica (surface area=300 m²/g; pore volume of1.65 cm³/g), and it is a material marketed under the trade name ofDavison 952 by the Davison Chemical Division of W. R. Grace and Company.

The alkylmagnesium compound is preferably a dialkylmagnesium having theempirical formula RMgR¹ where R and R¹ are the same or different C₂-C₁₂alkyl groups, preferably C₂-C₈ alkyl groups, more preferably C₄-C₈ alkylgroups, and most preferably both R and R¹ are butyl groups.Butylethylmagnesium, butyloctylmagnesium and dibutylmagnesium arepreferably used according to the present invention, dibutylmagnesiumbeing the most preferred.

The transition metal compound is preferably a titanium compound,preferably a tetravalent titanium compound. The most preferred titaniumcompound is titanium tetrachloride. Mixtures of such titanium metalcompounds may also be used.

The optional electron donor is preferably a silane compound, morepreferably a tetraalkyl orthosilicate having the formula Si(OR)₄ whereinR is preferably a C₁-C₆ alkyl compound. Typical examples of tetraalkylorthosilicate include tetramethoxysilane, tetraethoxysilane,tetraisopropoxysilane, tetrapropoxysilane, tetrabutoxysilane,tetraethoxysilane and tetrabutoxysilane being the two most preferredones.

The cocatalyst which can be used is preferably an organometalliccompound of a metal from groups I to III of the Periodic Classificationof the Elements, such as, for example, an organoaluminum compound, e.g.dimethylaluminiumchloride, trimethylaluminium, triisobutylaluminium ortriethylaluminium. Triethylaluminium is preferably used.

The catalyst can be used as it is or optionally in the form of a coatedcatalyst or prepolymer containing, for example, from 10⁻⁵ to 3,preferably from 10⁻³ to 10⁻¹, millimoles of titanium per gram ofpolymer. The process of the invention is particularly suited to the useof a non-prepolymerized catalyst, preferably to the direct introductionof a titanium magnesium silica supported catalyst.

The monohalogenated hydrocarbon compound of the present invention ispreferably added to the reactor in an amount such that the resultingcatalyst activity (gram of polymer per gram of transition metal perhour) presents an increase of at least 30%, preferably at least 50%,more preferably at least 80%, when compared with exactly the sameprocess conditions in the absence of said monohalogenated hydrocarboncompound.

According to a preferred embodiment of the present invention, themonohalogenated hydrocarbon compound is added to the reactor in anamount comprised between 0.1 to 40 moles of monohalogenated hydrocarboncompound per mole of transition metal of catalyst introduced into thereactor, preferably in a mole ratio comprised between 0.2 and 40,preferably 0.2 and 10, more preferably 0.25 and 5. Said mole ratio ormole of monohalogenated hydrocarbon compound per mole of transitionmetal catalyst can be measured by any appropriate method; for example,it can be measured through the measurement of the transition metalcontent of the polymer powder.

The monohalogenated hydrocarbon compound can be added continuously orintermittently to the reactor. In the continuous gas phasepolymerisation process according to the present invention, it ispreferred to add continuously the monohalogenated hydrocarbon compoundto the reactor. Sufficient monohalogenated hydrocarbon compound is addedto maintain its concentration at the desired level.

The following non-limiting examples illustrate the present invention.

EXAMPLES Example 1

a. Catalyst Preparation

The catalyst is a silica supported catalyst which is the same as onedisclosed in the comparative example 1 of WO99/05187 (1 mmol DBM/gsilica, 0.44 mmol TEOS/g silica, 1 mmol TiCl4/g silica)

b. Manufacture of a Copolymer of Ethylene and 1-Butene

The operation was carried out in a gas phase polymerisation reactorconsisting essentially of a vertical cylinder of 74 cm diameter and witha height of 6 m, with a disengagement chamber above it, fitted in itslower part with a fluidisation grid and a recycling conduit connectingthe top of the disengagement chamber to the lower part of the reactor,the recycling conduit being equipped with a cyclone, a heat exchanger, aseparator, a compressor and feed conduits for ethylene, for 1-butene,for hydrogen and for pentane. The reactor was also equipped with a feedconduit for catalyst and a conduit for drawing off copolymer.

The reactor contained a fluidised bed of particles of polymer beingformed, which had a height of 3.5 m and through which passed a stream ofreaction gas mixture, which had an upward velocity of 40 cm/s, anabsolute pressure of 2.1 MPa and a temperature of 105° C.

The reaction gas mixture comprised, by volume, 47.6% of ethylene, 0.4%of 1-butene, 16.3% of hydrogen, 9.6% of pentane and 26.1% of nitrogen.

The reactor was fed with catalyst prepared previously at a rate of 14g/h. It was fed separately with triethylaluminium at a continuous rateof 170 mmol/h. Furthermore, it was fed separately with n-butyl chlorideat a continuous rate of 25 mmol/h. The n-butyl chloride injection pointwas located in the gas-recycling conduit. The molar ratio of thequantities of n-butyl chloride and triethylaluminium introduced into thereactor was kept at 0.15.

Under these conditions a copolymer free from agglomerate was drawn offat a rate of 141 kg/h, which had a density (non annealed) of 0.959g/cm3, a titanium content of 3.4 ppm, a Melt-Index MI2.16 of 7.75 g/10minutes. The catalyst activity was 425 g of polymer per millimole oftitanium per hour of reaction and per 0.1 MPa of ethylene pressure.

Example 2

a. Catalyst Preparation

All manipulations were conducted under a nitrogen atmosphere. Into a 240L reactor was placed 12 kg of Ineos ES70 silica, which was previouslydried under a nitrogen purge at 500° C. during 5 hours. Hexane (60 L)was added to the silica. Dibutylmagnesium (16.2 mol) was added to thestirred slurry at 50.degree. C. and stirring was continued for one hour.Tetrabutyl orthosilicate (TBOS, 12 mol) was added to the slurry(50.degree. C.) and stirring was continued for two hour. TiCl4 (19.2mol) was added to the reactor (50.degree) and stirring was continued foran additional hour. Hexane was then removed by drying under vacuumconditions at about 40° C. Yield was 20.4 kg and the weight percent ofTi was 4.3.

b. Manufacture of a Copolymer of Ethylene and 1-Hexene

The operation was carried in a similar reactor as in example 1. Thecatalyst used was the one described above and the comonomer was1-hexene.

The fluidised bed of particles of polymer being formed had a height of5.5 m and the stream of reaction gas mixture passing through it had anupward velocity of 53 cm/s, an absolute pressure of 2.1 MPa and atemperature of 85° C.

The reaction gas mixture comprised, by volume, 19% of ethylene, 4.9% of1-hexene, 5.1% of hydrogen, and 71% of nitrogen.

The reactor was fed with catalyst prepared previously at a rate of 25g/h. It was fed separately with triethylaluminium at a rate of 280mmol/h. Furthermore, it was continuously fed with n-butyl chloride at arate of 14 mmol/h. The molar ratio of the quantities of n-butyl chlorideand triethylaluminium introduced into the reactor was kept at 0.05.

Under these conditions a copolymer free from agglomerate was drawn offat a rate of 163 kg/h, which had a density (non annealed) of 0.915g/cm3, a titanium content of 6.7 ppm, a Melt-Index MI2.16 of 1.0 g/10minutes. The catalyst activity was 455 g of polymer per millimole oftitanium per hour of reaction and per 0.1 MPa of ethylene pressure.

Example 3

a. Catalyst Preparation

The catalyst was prepared substantially according to example 1 of U.S.Pat. No. 6,140,264 except that tri-n-octyl aluminum was added to thecatalyst precursor after the catalyst precursor was contacted withconventional quantities of an electron donor.

b. Manufacture of a Copolymer of Ethylene and 1-Butene

The operation was carried in a similar reactor as in example 1. Thecatalyst used was the one described above and the comonomer was1-bbutene.

The fluidised bed of particles of polymer being formed had a height of 5m and the stream of reaction gas mixture passing through it had anupward velocity of 52 cm/s, an absolute pressure of 2.1 MPa and atemperature of 84° C.

The reaction gas mixture comprised, by volume, 47.6% of ethylene, 16.7%of 1-butene, 10.5% of hydrogen, 7.1% of pentane and 18.1% of nitrogen.

The reactor was fed with catalyst prepared previously at a rate of 34g/h. It was fed separately with triethylaluminium at a rate of 207mmol/h. Furthermore, it was continuously fed with n-butyl chloride at arate of 27 mmol/h. The molar ratio of the quantities of n-butyl chlorideand triethylaluminium introduced into the reactor was kept at 0.13.

Under these conditions a copolymer free from agglomerate was drawn offat a rate of 254 kg/h, which had a density (non annealed) of 0.918g/cm3, a titanium content of 0.62 ppm, a Melt-Index MI2.16 of 1.8 g/10minutes. The catalyst activity was 4140 g of polymer per millimole oftitanium per hour of reaction and per 0.1 MPa of ethylene pressure.

Example 4

a. Catalyst Preparation

The catalyst was prepared according the same protocol described inexample 3.

b. Manufacture of a Copolymer of Ethylene and 1-Hexene

The operation was carried in a similar reactor as in example 1. Thecatalyst used was the one described above and the comonomer was1-hexene.

The fluidised bed of particles of polymer being formed had a height of5.0 m and the stream of reaction gas mixture passing through it had anupward velocity of 52 cm/s, an absolute pressure of 2.1 MPa and atemperature of 88° C.

The reaction gas mixture comprised, by volume, 47.6% of ethylene, 6.5%of 1-hexene, 6.3% of hydrogen, and 39.6% of nitrogen.

The reactor was fed with catalyst prepared previously at a rate of 50g/h. It was fed separately with triethylaluminium at a rate of 260mmol/h. Furthermore, it was continuously fed with n-butyl chloride at arate of 39 mmol/h. The molar ratio of the quantities of n-butyl chlorideand triethylaluminium introduced into the reactor was kept at 0.15.

Under these conditions a copolymer free from agglomerate was drawn offat a rate of 307 kg/h, which had a density (non annealed) of 0.918g/cm3, a titanium content of 0.75 ppm, a Melt-Index MI2.16 of 1.05 g/10minutes. The catalyst activity was 3133 g of polymer per millimole oftitanium per hour of reaction and per 0.1 MPa of ethylene pressure.

1. Process for the gas-phase (co-)polymerization of olefins in afluidized bed reactor using a Ziegler-Natta type catalyst, said processcomprising the addition into the reactor of an organoaluminiumcocatalyst and of a monohalogenated hydrocarbon compound, a. wherein themolar ratio of the monohalogenated hydrocarbon compound to thecocatalyst is comprised between 0.02 and 0.2, b. wherein themonohalogenated hydrocarbon compound is added to the reactor in anamount comprised between 0.1 to 40 moles of monohalogenated hydrocarboncompound per mole of transition metal of catalyst introduced into thereactor, and c. wherein the monohalogenated hydrocarbon compound isn-butyl chloride.
 2. Process according to claim 1 wherein theZiegler-Natta type catalyst is a silica supported Ziegler-Nattacatalyst.
 3. Process according to claim 1 or 2, wherein the molar ratioof the monohalogenated hydrocarbon compound to the cocatalyst ismaintained constant throughout the polymerization.
 4. Process accordingto claim 1, wherein the sole or main olefin is either ethylene orpropylene, and the optional comonomer is selected from but-l-ene,pent-l-ene, hex-l-ene, 4-methylpent-l-ene and oct-l-ene.
 5. Processaccording to claim 1, wherein the monohalogenated hydrocarbon compoundis diluted in a diluent in an amount comprised between 0.001 and 2 moleof monhalogenated hydrocarbon compound per 1 of diluent.
 6. Processaccording to claim 1, wherein the monohalogenated hydrocarbon compoundis not added in admixture with the catalyst.
 7. Process according toclaim 1, wherein the catalyst is a non-prepolymerized catalyst. 8.Process according to claim 7, wherein the catalyst is a titaniummagnesium silica supported catalyst which is directly introduced intothe reactor.
 9. Process according to claim 1, wherein the cocatalyst iscomprised between 0.02 and 0.15.
 10. Process according to claim 1,wherein the amount in b. is between 0.2 and
 40. 11. Process according toclaim 10, wherein the amount in b. is between 0.2 and
 10. 12. Processaccording to claim 11, wherein the amount in b. is between 0.25 and 5.13. Process according to claim 5, wherein the diluent is selected fromthe group consisting of butane, pentane and hexane.